Hydrocarbon conversion



(kt-31, 1944. E. L D ouvlLLE ETAL HYDROCARBYON CONVERSION i Filed oct.(15, 1940 Patented ct. 31, 1944 UNITED STATES PATENT o1=1=1cEIIYDROCARBON CONVERSION i Edmond l..A dDuville and Bernard L. Evering,

Chicago, lli.. anaignors to Standard Oil Com- NIW. Chicalo. lll..corporation oi' Indiana Application October l5. 1940, Serial No. 361,270

(Cl. 19d-48) 3 Claims.

This invention relates to certain new and useful improvements in theconversion of low antiknock hydrocarbon fractions to useful productssuch as high antlknock motor fuels and high antiknock aviation fuel. Theinvention has in view the provision of a combination process in whichthe low antiknock components of virgin petroleum naphthas aie treated inan integrated process which is highly efilcient in attaining theultimate object of maximum yield of high antiknock naphthas.

Virgin naphthas of motor fuel distillation range contain variableproportions of straight- 'chain 'parailln hydrocarbons and naphthcnichydrocarbons with minor proportions of aromatic hydrocarbons. Forexample. virgin naphthas from Pennsylvania crudes and Michigan crudesmay contain as high as eighty to ninety per cent parafiins, the majorpart of which are straightchain parafllns and as low as five to fifteenper cent of naphthenic hydrocarbons. On the other hand, naphthas fromMid-Continent crude and from California crude contain as high as twentyper cent to thirty-live per cent of naphthenic hydrocarbons and fromfifty to seventy per cent of paraffin hydrocarbons. Since thestraight-chain paraiiin hydrocarbons and the naphthenic hydrocarbons areof low antiknock value it is highly desirable to convert the former tobranched chain parafi'in hydrocarbons and the latter to aromatichydrocarbons boiling within the gasoline boiling range in order toobtain maximum yields of high antiknock value naphthas. i

It is an vobject of this invention to provide an improved process forconverting the straightchain parafiins in a parainic naphtha which alsocontains naphthenes and aromatics, to isobutane. Another object of thisinvention is to convert the straight-chain parain hydrocarbons of such anaphtha to condensible oleiins. Still another object of this inventionis to react the isobutane and the condensible olens obtained from thestraight-chain paraiins to form high antiknock branched chain parafnicnaphthas which distill 'in the aviation gasoline boiling range. Afurther object of this invention is to raise the' antiknock propertiesof a low antiknock naphtha by removing therefrom, in a catalyticconversion process, a substantial part of the low antiknockstraight-chain paraffin hydrocarbons. Another object of this inventionis to convert the low naphthenic hydrocarbon components of a paraiiinicnaphtha to high antiknock aromatics. A still further object of thisinvention is to provide a continuous process in which the catalystsemployed may'be utilized eiectively for the production of high antiknockmotor fuels v metallic oxide activated alumina catalyst.

the accompanying drawing which is a diagram-v matic representation of anapparatus for carrying out our invention.

We have found that straight-chain paraffnic naphthas can becatalytically converted to a gaseous product which is largely isobutaneby passing the naphthaover a gel catalyst of-the alumina activatedsilica type at low space velocities o! the order of 0.01 to 1.0 volumeof liquid per volume of gross catalyst space per hour and attemperatures withinV the range-of 575 F. to 850 F. This conversionprocess is carried out at absolute pressures within the range of fifteenpounds and sixty-live pounds per square inch. Moreover. we have foundthat by operating in the upper part of the above temperature range, say800 F. to 850 F., the gaseous product consists primarily of isobutaneand condensible oleilns in such proportion that the gaseous mixture maybe alkylated to produce bz anched chain paraiiinic hydrocarbons boilingwithin the gasoline range. We have also found that if a paraflinicnaphtha is passed over a solid catalyst comprising a mix2 ture of acracking. type catalyst, such as alumina activated, silica and adehydrogenation type catalyst, such as metallic oxide activated alumina,at temperatures within the range of 750 F. to 900 F., at spacevelocities within the range of 0.01 and 1.0 volume of liquid feed pervolume of gross catalyst space per hour and at absolute pressures withinthe range of ten pounds and sixty pounds per square inch, relativelylarge volumes of condensible olens are produced. Under these operatingconditions,any naphthenes present in the feed to such a catalyst bed arealso dehydrogenated to high antiknock aromatic hydrocarbons.

In practicing the invention, we prefer to fractionate at least a majorpart of the naphtha feed stock in order that we may convert the paraffin-hydrocarbons of less than seven carbon atoms directly to the branchedchain paralns. This lower boiling fraction of the naphtha feed stockwill contain only limited amounts of aromatic hydrocarbons an'd hence Weare able to isomerize it to branched chain paraffin hydrocarbons overaluminum halide type catalyst. Aromatic hydrocarbons tend to deterioratethe aluminum halide catalyst. Thus we reserve the C1 and higher fractionof the feed, which contains aromatics and naphthenes in addition toparaiilns, for conversion over the alumina activated silica gel crackingcatalyst and/or for conversion over the mixed alumina activated silicagel catalyst and The isomerized Cs and C5 hydrocarbons are blended withthe alkylated isobutane and utilized as aviation fuel or the blend maybe added to the converted naphtha to produce high antiknock motor fuel.The C4 and lighter fraction of the virgin naphtha feed passes to thealkylation step which C4 and lighter fraction, a Cs-l-Cs fraction and aC1 and higher fraction. The Cv and higher fraction is heated to atemperature within the range of 575 F. to 850 F., preferably 750F. to800 F., and is thenfpassed over a catalyst bed in a conversion zonewhich may be designated as a cracking zone in which 'the catalystcomprises an alu- The product from thisA mina activated silica gel.conversion step is fractionated into an overhead gaseous fraction whichis extremely high in isobutane content. 'I'he major part of the liquidfraction is recycled to the heating step' and thence to the saidcracking zone to convert more of the straight-chain paraiiins toisobutane. A part of the liquid fraction is returned to the furnace andreheated in a separate coil to a higher temperature, for example, 750`F. to 900 F., but preferably from 850 F. to 900 F., and then passed to aconversion zone containing a mixture of .a cracking catalyst and adehydrogenation catalyst such as alumina activated silica mixed with ametallic oxide activated alumina catalyst. We have found that chromiumoxide activated alumina is a suitable dehydrogenation catalyst formixing with the alumina-silica catalyst in this conversion Zone whichmay be designated as a dehydrogenation zone` The product from thisconversion step is fractionated and the high antiknock value liquidfraction from which a major part of the paramns have been removed,either by cracking to isobutane in the 'cracking zone or to condensibleolefins in the dehydrogenation zone and in which the naphthenes havebeen converted to aromatics, is withdrawn for blending. y If desired theCv and higher fraction may be sent to the dehydrogenation zone and apart of the liquid fraction from the fractionation of thedehydrogenation product may be recycled to the cracking zone, in whichcase the products from the two conversion zones may befractionated in acommon fractionator, thus eliminating the necessity of a separatefractionation of the catalytically cracked product. The isobutane plusnormal butane overhead product from the fractionation of the product ofthe cracking step is alkylated with the condensible oleiins fromthefractionation of the product of the dehydrogenation step byintimately contacting the liquefied mixture with sulfuric acid of 95% to103% concentration at temperatures chain C5 and Ce paramn hydrocarbonsby intimately contacting this fraction and the normal butane -withaluminum halide or aluminum halide-hydrocarbon complex catalyst,preferably aluminum chloride or aluminum chloride hydrocarbon complex,at temperatures within the range of 250 F. to 400 F. The amount ofcatalyst used may be from ve to fty per cent of the Weight of thehydrocarbon feed. The catalyst is activated by the addition of hydrogenchloride to the reactor to the extent of from five-tenths to five percent of the weight of the hydrocarbon feed. The isomerization reactionis carried out in the presence of hydrogen or a gas containing hydrogen.We prefer to use at least a part of the hydrogen and methane containinggas from the dehydrogenation step for this purpose. Pressures in theisom.

erization step may be within the range of 200 to 3000 pounds per squareinch. The product from the isomerization step is passed to thedehydrogenation product fractionating step from which the isomerizedC5+Cs cut is withdrawn as a high antiknock blending naphtha, isobutanebeing separately recycled to the alkylation step.

In converting low antiknock value naphthas, which contain very lowpercentages of naphthenes, to high antiknock blending naphthas by ourprocess it is sometimes unnecessary to utilize the dehydrogenation stepin which the feed is contacted with the mixed cracking anddehydrogenation catalysts. We have found (see Table I below) thatwhenoperating the cracking reactor at a temperature of 823 F. thegaseous product contains a weight ratio of isobutane to total gaseousolefins reactive in a sulfuric acid alkylation process of approximately29.0 to 27.6. which corresponds to a mol ratio of approximately 22.0 to25.6. Keeping in mind that one mol of olefin reacts with one mol ofisobutane in the alkylation reaction, it is readily seen that within therange of 30 to 150 F. for periods of from five minutes to sixty minutesand at pressures sucient to keep the reactants in the liquid phase. Theratio of isobutane to condensible olens should be at least 2 mols ofisobutane for one mol of oleiins. The alkylated product is separatedfrom the unreacted gases by fractionation and used as a high antiknockvalue blending naphtha for aviation gasoline or it may be blended withthe high antiknock value naphtha, from the cracking and dehydrogenationsteps. Unreacted normal butane is passed to a catalytic isomerizationstep to produce additional isobutane for the alkylation reaction.

The Cs-i-Cs fraction from the virgin naphtha, feed is isomerized alongwith the recycle normal' butane from the alkylation reaction to producesimultaneously isobutane and branchedy by isomerizing the normal butaneproduced in the` cracking reactor and also the butane' in the overheadfraction of the fresh feed it is possible to maintain a properlyproportioned feed to the alkylation reaction without using thedehydrogenation reactor to produce olens. -Any deficiency in hydrogenfor pressuring the isomerization reaction may be furnished from anoutside source for this type of operation. The nal products from thistype of operation are the same as from the process when both thecracking and dehydrogenation reactors are used except lthat the productcontainsL a lower percentage of aromatic hydrocarbons. This type ofoperation may also be used when it becomes necessary to regenerate thecatalyst in the dehydrogenation reactor.

` TABLE I CA'rALY'rIc CONVERSION or NORMAL HEPTANE CUT Space velocity0.06 vol. of liquid feed per vol. gross Temperature F 76H SIR x03 Weightper cent reacted l. 4 3. 7 I4. 0 Product-Weight per cent:

H i. 4. 2 4. 8 4. 4 2. 3 4. li

l. l l. 6 2. 7

I3. 8 ll. 6 l0. 2

I8. 3 I7. 2 l5. 5

NC4Hm 7. 2 9. 0' 5. -l

' scribed below.

In a third embodiment of the invention the dehydrogenatlon reactor maybe operated without the cooperation of the cracking reactor. Referringto Table II we have found that if a paraftln hydrocarbon fraction ispassed over a mixed catalyst comprising an alumina activated silica geland a chromium oxide activated alumina catmal butane is formed but themajor part of the gaseous product consists of condensible olefins' whichmay be polymerized by either the cold sulfuric acid process or the hotsulfuric acid process. The polymerization process can be carried out inthe alkylation reactor and the spent alkylation sulfuric acid catalystcan be reduced in concentration to a concentration within the range of60% to 80% acid and used for this step.

The polymerized product4 can be fractionated to' The polymerization 4bythe hot acid method isA alyst at 893 F.' no isobutane and very littlenorcarried out at temperatures within the range from about 160 F. to 200F. at atmospheric pressure with contact time within the range of fromthirty seconds to three hundred seconds.

'I'his type of operation is desirable when the feed stock consists of anaphtha containing a rela- I tively low percentage of paramos and a highpercentage of naphthenes.; This type of operation may also be utilizedduring those periods when it becomes necessary to regenerate thecatalyst in the cracking reactor. During operation according to thismethod the isomerizer may be used to isomerize only the Cs-l-Cs fractionof the fresh feed.

VTABLE -II CA'rALYrrc Conversion or NORMAL `Hari-Am: CU'r Space velocity0.06 vol. of liquid gross catalyst space per hour. mosphericCATALYST-MIXED SILICA-ALUMINA ANT) 95% A1203.5% Creo;

feed per vol. of

Pressure-at- Other embodiments of our invention, such as passing freshfeed to the cracking reactor and to the dehydrogenation reactor inparallel with separate fractionation of the product and separate recycleto the respective reactors or with fractionation in a common zone andconsequent recycle of mixed liquids to either or both reactors or acombination of any of themethods of operation described above arecontemplated by this invention. For the purpose of illustrating theinvention operation ,with a feed stock such as a Mid-Continent naphthacontaining straightchain parafns, aromatics and naphthenes is de-Referring now to the drawing, which is a simpliiled ilow diagramillustrating one embodiment of our invention, a virgin naphtha boilingwithin the range from about 60 F. to about 400 F. is

the fresh feed through fractionator I6 in order to separate the maiorpart of the lower boiling fraction of the feed which may beisomerizeddirectly to high antiknock value light naphtha suitable for blending inaviation gasoline. Fractionator I6 is provided with reux means andreboiler means I8. A side stream in line |33 containing primarilypentanes and hexanes is withdrawn from fractionator I6 by pump |34 asfeed stock -for the isomerization step which is described in more detailbelow. The C4 and lighter overhead fraction from fractionator i6 ispassed by,valved line 96 to alkylator |05 which is also described below.

The C7 and higher fraction of the naphtha is passed by means of pump 20in line |9 to valved line 2| where it is mixed With non-fractionatedvirgin naphtha feed and the mixture is passed through valve 23 in line22 ing coil 25 in furnace 24, valve 90 in line 89 being closed. Thenaphtha is heated in coil 25 temperature within the range of 575 F.,preferably '750 F. to 800 temperature range it is passed 26 and valve 21to lcatalyst chamber 28 which is lled with a refractive gel catalyst ofthe alumina activated silica type. This catalyst may F. to 850 F.,within which be prepared by immersing commercial silica gel in anaqueous solution of an aluminum salt. draining ofi the supernatantsolution, Washing the gel, and drying and heating the product forseveral hours at temperatures of from 800 F'. to '1100 F. The finishedcatalyst contains from 95% to 99.5% silica gel activated with from 0.5%to 5% alumina. We may also'employ an alumina-activatedv silica gelcatalyst prepared by cogelling an alumina sol with a silica sol in suchproportion as to produce a cogel comprising from about one to twentyparts by weight of alumina and from about to 99 parts by weight ofsilica in the dried gel product. Other refractory type crackingcatalysts such as acid treated clay and silica gel activated with suchmetallic oxides as magnesium oxide and boron oxide may also be usedincracking chamber 28. The space velocity in catalyst chamber 28 atabsolute pressures within the range of l5 pounds per square inch and 65pounds per square inch, is Within the range of 0.01 and 1.0 volume ofliquid feed per volume of apparent catalyst space per hour based on theparain hydrocarbon content of the feed. We prefer to operate `at spacevelocities within the range of 0.04 and 0.1 volume of paraiinhydrocarbons per volume of catalyst per ho-ur or at such a spacevelocity that from about 5% to about 15% of the straightchain parafiinhydrocarbons in said naphtha are converted per pass through the reactionzone.

. The product from catalyst chamber 28 is with.

and it is passed which leads to heattoa` via transfer line which isoperatedA product from reaction chamber 28 the product may be passeddirectly to fractionator ll via valved lines i2 and 53 with valve iidopen and valve l5 closed. Fractionator lll with reflux producing meansi6 and reboiler lil serves to separate the C4 and lighter -hydrocarbonswhich contain a large percentage of isobutane from the liquid product.The butanes pass via valve 95 in line 95 and line Tl to the alkylatorfeed line 98. The liquid product is withdrawn from' fractionator ilthrough side drawoi line 55. Hydrocarbons boiling above normal motorfuel boiling range may be drawn off through bottom drawoi 58a. By properadjustment of valve 50 in line 59 which joins line 58, valve 5i in line52, valve 55 in recycle line 55 and valve 55 in nished product line 5l,the liquid product from fractionator lll may be sent in part to coil 5lin furnace 25 through lines t5 and 05 and thence by linel 92 todehydrogenation reactor 50, and the major part of the product may berecycled by means of pump 53 through lines 52 and 55 to coi-l 25 infurnace 25 and thence to' reaction chamber 20- for the production ofadditional isobutane. 'We prefer to operate with4 valve 55 in line 5lclosed when processing a vir-gin naphtha containing appreciable amountsof naphthenes in order that these be converted to aromatics in reactor55 before blending the liquid product to form high antiknock gasoline.

Referring to Table I above, a conversion per pass of approximatelyeleven per cent by weight of-the paran hydrocarbons may be expected inreactor 20 when operating within the range of 750 F. to 800 F. In orderto convert a major part of the paraln 'hydrocarbon components of thefeed in reactor 28 we maintain a recycle ratio of from about eight toabout twenty parts by weight of recycle to one part by weight of virginfeed. When processing feed stocks containing a relatively highpercentage of paraffin hydrocarbons the recycle ratio will be intheupper part of the above range while a feed stock containing a relativelylow percentage of paraln hydrocarbons will not necessitate such highrecycle ratios.

In order to more completely reduce the paraifin hydrocarbon content oftheproduct in .side drawoff line 58 and thereby produce condensibleolefins and also to dehydrogenate the naphthenes-therein contained toaromatics, valve 50 is partially opened and the stream is directed tofurnace coil @l in furnace 20 via lines 85 and 89 by means of pump 53.This stream is heated to a temperature within the range from about '750P'. to about 950 F., preferably 825 F. to 900 F., and passes by line 92to catalytic reaction chamber 58 which is filled with a mixed catalystcomprising from about 1 to 10 partys by weight of alumina activatedsilica .gel catalyst intimately mixed with from about 99 to 90 parts byweight of a. metal oxide-activated alumina catalyst such as an activatedalumina supported chromium oxide catalyst. We prefer to use in reactor58 a catalyst comprising alumina impregnated with 5 to- 20 parts byweight of chromic oxide, preferably parts by Weight of chromic oxide,mixed with a second catalyst prepared by absrptolytically depositingA1203 on Si02 gel as described above. The chromic oxideactivated aluminacatalyst may be prepared by immersing commercially activated alumina inan aqueous solution containing the desired amount of chromic oxide inthe form of the acid or suitable salt, evaporating oi the water, dryingthe product and heating the same at 1200 F. for one hour. The powder isthen pelleted to the desired particle size. Other metallic oxides. suchas molybdenum oxide and vanadium oxide may also be used to activate thealumina by using the appropriate water soluble acid and/or water solublesalt. The space velocity in reactor 50 may be within the range of 0.01to 1.0 volume of paraffin hydrocarbons (liquid) per gross volume ofcatalyst per hour. The operating pressure (absolute) in reactor 58 iswithin the range of from 10 pounds to 65 pounds per square inch. Inreactor 55 the major part of the remaining paraffin hydrocarboncomponents of the naphtha are dehydrogenated and cracked to a gaseousproduct, a high weight per cent of Which product consists of condensibleolens suitable for feed to the alkylation step which is described below.We have found (Table II above)r that a conversion of 14 per cent perpass of pararn hydrocarborm over the above mixed catalyst may beobtained at a temperature of 893 F., at space velocity of 0.06 volume ofliquid feed per volume of catalyst space per hour and atmosphericpressure and hence, in order to reduce the parainic content of theresidue naphtha to a minimum, we prefer to operate at a recycle ratio offrom about five to twelve parts by weight of naphtha bottoms from thefractionation of the product from reactor 58 (fractionator l2) toone'part by weight of feed to reactor55 from line (i9. The naph- `thenesof4 the naphthaare converted to aromatics in reactor 58 thus furtherincreasing the antiknocl: value of the residual naphtha.

The product from reactor 50 passes via line 59, valve 00, cooler 5i,compressor C32 and line 03 to hydrogen release drum 55 where hydrogenand non-condensible gases, such as methane, ethylene and ethane areseparated and pass via lines-55 and 6l to the isomerization reactorSill. These may be passed in part to the fuel burning line throughvalved line 55 which joins line G5 by adjusting valve 55 in line 55 andthe valve in line 58, or any excess of these gases not required inisomerization reactor Hill, when reactor 50 is on stream in the reactionpart of the cycle, may be passed to storage for use in reactor iS' whenthe catalyst in reactor 58 is being regenerated. Still another use forthis hydrogen is for the hydrogenation of polymer produced whenalkylator |05 is being used as a polymerizer. which operation isdiscussed below. Also, the hydrogen may be used to suppress carbonformation in reactor 58. Condensate from drum 50 to which product fromline 53 is introduced at pressures within the range of 300 pounds to 400pounds per square inch.' passes by line l0 and pump 'H to fractionator72 which is provided with reboiler i3 and reflux means lil. Residualnon-ccndensible gases pass overhead via valved linel'i5 to the fuelburning line. A side stream, predominantly C3 and C4 hydrocarbons, iswithdrawn from fractionator 'l2 through line 16 and passes via line Tlto alkylator hydrocarbon feed line 98. This stream in line 16, whichconsists in largepart of condensible olelins, may be directed throughline '18 when alkylator |05 is being used as a polymerization reactionchamber. Fractionator 'l2 is also used to separate the liquid isomerizerreaction products which are introduced to fractionator via line |55 andhence a Cs-l-C side stream is trapped out of fractionator 12 which ispassed via lines 79 and l30 by pump to product line 51 for blending withthe aromatic bottom product from fractionator 'l2 or this Cs-l-Cc streammay be passed via, lines 18 and 18a to an aviation gasoline blendingtank (not shown). The aromatic naphtha bottoms are withdrawnfrom-fractionator 12 via line 8| and pump 82, a major part of thisstream being .recycled to coil 8| in furnace 24 vialine 85, theVremainder being eliminated as high antiknock naphtha via line 51. anyhydrocarbons boiling above normal motor fuel boiling range may bewithdrawn through bottom drawoff 8|a.

As stated above, we may operate our process using a common tower tofractionate the prod- 'ucts from the catalytic cracking step and thecatalytic dehydrogenation step. By properly adjusting valve 23 in line22 and opening valve .88 in line 88 the C1 and higher fraction may bepassedl via heating coil 8|, line 92 and valve 83 to dehydrogenator 58for the initial conversion Y step in which condensible oleflns andhydrogen are produced. The product is flashed in release drum 64 andfractionated in fractionator 12 and the major part of the liquid bottomfraction may be recycled via line 8|, pump 82, line 85 and lline 86 tocoil 25 via line 26 and thence vto reactor 28 where isobutane isproduced in the cracking operation. The product from reactor 28 iswithdrawn via line 28 and passes to the dehydrogenator product line 58.By adjusting valve 84 in line 8|, valves 83 and 88 in line 85 and valve81 in line 88, the liquid fraction from fractionator 12 may beproportioned into nished high antiknock fraction drawoif, recycle feedto the dehydrogenation step and recycle feed to the catalytic crackingstep as desired. Virgin C7 and higher naphtha may be passed in parallelto reactors 28 and 58 by partially opening valve 23 in line 22. Thechief advantage of passing the major portion of the virgin feed tothedehydrogenation step with operation of the catalytic cracking reactorprimarily on recycle from the dehydrogenation step is possible heateconomy without sacrice of the desirable dehydrogenation of naphthenesto aromatics. 'Ihis type of operation-is desirable when the virgin feedstock is high in naphthenic hydrocarbon content. When the virgin feedstock is highly paraflinic with very low naphthenic hydrocarbons presentin the feed, reactor 28 may beV operated at higher temperatures, say 800F. to 850 F., and the product may be fractionated either in fractionator4| or 12 as described above.

Referring now to the alkylation'step of our process, the butane overheadstream from fractionator I8 passes via line 88 and valve 81 as alkylatorfeed stream to line 88 where it is mixed with the isobutane rich gaseousstream from line 11 which connects with fractionator 4| overheaddischarge line 85. Isobutane may be introduced to line 88 from anexternal source through valve line 88. condensible oleflns pass viafractionator 12 drawoi line 18 to line 11 and thence to line 88 or theseoleilns may be added directly to the acid catalyst valved line |86through line 18. Condensible olefins from an,external source may beadded either directly to the acid catalyst by means of lines |88 and |8|or the olens may be' `passed through lines |88 and |82 to line 11. Themixed butane and olefin stream is compressed by compressor |83 andpasses by line |84 to alkylation reactor |85 where the mixture isintimately contacted with sulfuric acid of 95% to 103% strength which isintroduced to line |84 via line |86 and valve |81 by means of pump |88.Alkylation reactor |85 is equipped with a mechanical stirrer |88 and atemperature control jacket ||8 15 with contact times withink the rangeof from about five -minutes to sixty minutes. The alkylate-acid mixtureis transferred to settler by means of pump ||4 in valved line H3. Insettler-*H5 the acidseparates as a lower layer and it may be recycledthrough line H6, pump H1, line H8 and line |88 to line |84. Spentsulfuric acid may be withdrawn through valved line I8 to be eitherconcentrated for reuse in the alkylation step or to be diluted for usein the polymerization step described below. The alkylate and -unreactedgases are passed vialine |28 and pump |2| to a caustic wash step and awater wash step, which steps are not shown in the flow diagram. Thewashed product is heated in heater |22 and passes by line |23 tofractionator 24 which is equipped with reboiler |25 and reflux producingmeans |26. Propane and lighter hy- .drocarbons pass from fractionator|24 through valved line |21 to the fuel burning line. A side streamconsisting predominantly of unreacted isobutane and normal butane iswithdrawn through valved line |28. This stream may be recycled to thealkylationstep feed line 88 via valved lines |28 and |28 or toisomerizer feed line |33 via line |28 for the production of additionalisobutane. A second side stream consisting of alkylate having an octanenumber of to CFR-M and an end point suitable for blending in aviationgasoline orsuitable for blending in finished motor fuel is withdrawnthrough line |38 and passes to line |3| and thence to line 18a forblending with the isomerized Cs-i-Cs cut or the alkylate may be sentdirectly to finished motor fuel line 51 via line |38. Bottom drawoiffrom fractionator |24 consists of a small amount of alkylate distillingabove the gasoline boiling range when operating this fractionator toproduce 400 F. end point alkylate drawoff in line |88 for blending withmotor fuel. When so operating these bottoms are withdrawn through valvedline |38a and accumulated as additional feed to alkylator |85 which maybe operated at higher temperatures, for example, from F. to F., as adepolymerization-alkylation reactor to produce a product boiling withinthe gasoline range when charging such a stock as the bottoms from tower|24 with condensible olens. When operating to produce an alkylate forblending in aviation gasoline thehigher -boiling fraction may be furtherfractionated to produce additional alkylate of 400F. maximum boilingpoint for blending in motor fuel and the fraction distilling above 400F. may be accumulated and used as described above.

As stated above we prefer to isomeriz'e the C5+Cs` fraction of thevirgin naphtha directly to highly branched parafns for`blending in thefinished products. On the other hand, it is not desirable to crack thisfraction in reactor 28 since the increased yield of non-condensiblegases which represent a volume loss of product would be increased andmoreover this light fraction is essential in blending to form a balancedaviation gasoline or motor fuel product to furnish desirable volatilitycharacteristics and if isomerized the fraction contributes materially tothe antiknock value of such products. Hence this frac tion is drawn oilfrom fractionator I8 through through which may be circulated a suitablecoolline |33 by pump |30 along with recycle butane from line |28. Alight naphtha fraction containing C4, C5 and Cs paraflins may also beintroduced from an external source via line |3341. The mixed hydrocarbonstream is heated in heater |35 which may be a heat exchanger in whichthe heat is furnished by the product from reactor 20 or reactor 58.Aluminum chloride or aluminum chloride-hydrocarbon complex catalyst,which is` activated with hydrogen chloride, isintroduced to the hotstream in line |36 via valved line Mil and pump M2. Hydrogen chloride isadded to the catalyst in line Uil through valved line 850 and themixture passes tc isomerizer itl which is equipped with rapid stirringmeans i323. Reactor lill may be suitably jacketed and a heating uid iscirculated therethrough by means of valved connecting lines i353 andi130. The isomerization reaction is carried out in the presence ofhydrogen or hydrogen mixed with non-condensible gases supplied by thedehydrogenation product via line til and compressor M3. Hydrogen may befurnished from an external source through valved line and line l. Weoperate isomerization reactor itil Within the range of 250 F. to 400 F.,preferably at about 330 F., and at pressures Within the range of 200pounds to 3000 pounds per square inch. The contact time in reactor ilmay be :from to 100 minutes.

Part of the catalyst plus isomerized product mixture is withdrawn fromreactor l'l by pump |05 through line lllrand passesl by line it toseparator lill which operates as a catalyst settling Zone and also as arelease drum in which the hydrogen and hydrogen chloride is separatedfrom the liquid product. The pressure in separator Mill is lowered tothe range of 150 pounds to 350 pounds per square inch by openingpressure release valve |53 in line |52 and the mixed hydrogen andhydrogen chloride is recycled via line |50 to catalyst feed line Ml andthence to reactor |3l 'through pump |02 and line |36. The aluminumchloride catalyst settles from the liquid hydrocarbons in separater li'lfrom which it is withdrawn by pump |69 in line |00 and it is recycled toreactor itl through lines ll, lill and |36. Spent catalyst may bewithdrawn from the cycle by means of valved line |50. The hydrocarbonmixture comprising butane, isobutane and the isomerized Cs-l-Ce fractionpasses via valved line |55 to line l0 which leads to fractionator l2from which the butanes are recycled to the alkylation step and theisomerized Ct+Ce fraction of approximately 82 octane number (CFR-M)passes therefrom as blending naphtha to be disposed of as describedabove. If desired the hydrocarbon stream in line |55 may be subjected toa hydrogen chloride absorbent or neutralizing agent such as solid sodiumhydroxide or other neutralizing material in order to remove the lasttraces of hydrogen chloride before introduction of the stream to linel.

Referring now to the regeneration of the catalysts in reactors 28 and58, these may be regenerated by oxidation of the carbonaceous materialdeposited thereby means of an oxygen containing mixture of gases such asair which may be diluted with nue gas to reduce the oxygen percentage ofthe gas. Although we have described catalyst chambers as singlereactors, `these may be provided as multiple units to be operated inparallel banks, either continuously or intermittently, the catalyst inone or more of the chambers of the multiple units being regeneratedwith'- out interruption of the continuous process. Re-

actors |05 and |31 may also be in multiple. However, the regeneration ofthe catalyst in reactors 23 and 58 may be accomplished without the useof multiple chambers by a slight change in the mode of operation when itbecomes neces- Ysary to regenerate either of these catalyst beds.

For example, when the catalyst in reactor 20 becomes spent we mayisolate reactor 28 by closing valves 2l and 30 and valves |57 and |59are opened'to admit and discharge inert gas from reactor 28 throughlines l5@ and |58 to remove hydrocarbon vapors. After the reactor ispurged the inert gas is gradually enriched 'with oxygen or air and thecarbonaceous material is removed as an oxidation product. The reactor isagain purged with inert gas, after which reactor 28 is in condition foruse in the hydrocarbon conversion process. During the regeneration ofthe catalyst in reactor 20 all of the virgin feed may be fractionated inractionator iii in order to obtain maximum isobutane in the process. Thesupply of lsobutane for alkylation during this catalyst regenerationperiod may be augmented by introducing isobutane from an external sourcevia valved line il@ or by introducing normal butane or a normalbutane-isobutane stream to butane isomerizer i3? via valved line 20@ andlines l2@ and i3d. As an alternate method or `operation during theregeneration of the catalyst in reactor 28 we may operate alkylator itas a polymerizer using diluted spent allrylation acid in the hot acidprocess vto polymerize the olefins produced in. reactor Reactor 605 isoperated at temperatures within the range of 165 F. to 212 F., the acidstrength being Within the range of 60% to 80% sulfuric acid whenpolymerizing oleiins. rl`he polymer may be separated from the acid insettler liti and separated from unreacted gases in fractionator M0,polymerof suitable boiling range being with- 'drawn via line |30 forblending in motor fuel ork of suitable boiling range for hydrogenationover suitable catalyst and under conditions well known in the art, theproduct to be blended in aviation gasoline.

When it becomes necessary to regenerate the catalyst in reactor thereactor may be isolated by closing valves 03 and 60 in lines 02 and 59,respectively, and inert gas and regenerating gas is admitted via valvedline i to reactor 08 and exits via valved line |02, valves 46| and M53belng open. During the regeneration of the catalyst in reactor 58reactor 20 may be operated at temperatures Within the range of about 800F. to 850 F. within which temperature range sufflcient condensible olensare produced along with lsobutane to alkylate the isobutane in reactorIll. Any make-up hydrogen required for the isomerization reaction inreactor itl during the regeneration of catalyst bed 5&3 is furnishedfrom storage via valved line 69, which leads `to lines G5 and 51 andthence to reactor |311. During the regeneration of the catalyst inreactor 00 the product from reactor 28 may be sent to fractionator 0|,valve 50 in line 49 being closed and verted in reactor 58.

. 2,861,611 bimng the recycleboaoms with fresh feed in line 22.

We may also operate reactors28 and 58 in parallel as well as in seriesas described above. For example, virgin naphtha of wide boiling range orthe C1 and higher fraction of virgin naphtha may be delivered to lines22 and 89 from line 2l simultaneously and mixed therein with recyclefrom fractionators 4I and 'l2 by lines 55 and 85 respectively, thearomatic bottoms eliminated from each fractionator beingdischarged tofinished product line -57 or the products from reactors 28 and 58 may besent to a common frac-1 tionation system by passing the product fromreactor 28 through line 29 to reactor 58 product line 59 viavalves 42and 45, valves 3i, 36, 40, 50

the optimum amounts of isobutane and condensible olens.

By means of our process a typical Mid-Continent straight-run naphthaboiling between 60 F. and 400 F. and having an octane number of 45 to 50can be converted into high grade motor fuel` of 85 to 95 octane numberin yields of 80 to 90% on a, volume basis. The Cs-i-C cut of such anaphtha will generally be about to by volume and will have an octanenumber of 65 to 70. This cut is converted, at least partially, toisobutane in reactor itl, while the residue is isomerized to an octanenumber of 80 to 85 on a butane-free basis.. The volume recovery of thismaterial, including the isobutane, is 95% to 105% by volume.

The C1 and higher cut of the naphtha generally contains 15 to 25%aromatics having an average octane number of 100, to 40% naphtheneshaving an average octane number of from to 65 and 40 to 50%k parafilnshaving an average octane number of zero to l5. Under normal operatingconditions about to 70% of the paraiiins are vconverted togases inreactor 28 while to 90% of the remaining paralns are con- As previouslydescribed these gases are principally Ca 4and C4 hydrocarbons and areconverted in reactor |05 to an alkylate of to 95 octane number in yieldsof 65 to 75% on the basis of the paramns converted in thecondensiblegases may be separated from the l conversionA product andrecombined to form naphthas having highantiknock value which can beblended baci:V with the residue naphtha to make'motor fuels of greatlyincreased antilcnock value. Although we have described 'our invention indetail and, therefore. utilized certain speciil'c terms and languageherein, itis to be understood that the present disclosure isillustrative rather than restrictive and that changes and modificationsmay be resorted to without departing from the spirit or the scope of theclaims appended hereto.

We claim:

1. The process oi.' producing isobutane which comprises contacting apetroleum naphtha with a solid, silica-alumina cracking catalyst in areaction zone at a temperature above 575 F. but below 850 F., at apressure within the approximate range of about 15 to 65 pounds persquare inch and at a low space velocity of the order of 0.01o to 1volume -of liquid naphtha per hour per volume of catalyst space in thereaction zone, removing the products from said reaction zone,fractionating the products to obtain a fraction rich in isobutane and afraction containing hydrocarbons of the naphtha boiling range andrecycling at least a substantial part of the lastnamed fraction to saidreaction zone for obtaining further production of isobutane.

2. The process of claim 1 wherein the petroleum naphtha is rich instraight-chain parailin hydrocarbons and wherein the space velocity issuch that from about 5% to about 15% of the straightchain paraillnhydrocarbons in said naphtha is converted to normally gaseous productsper pass through said reaction zone.

reactors 28 and 58. The naphthenes are substantially converted to4aromatics in reactor 58 so that the liquid product from the bottom offractionator 12 will be highly aromatic and have an octane number of to100. As hereinbefore described these high octane number products can beblended together or used separately for special purposes.

We have described a process whereby the low 3. The process oisimultaneously producing isobutane and a motor fuel fraction having ahigher antiknock value from a low antiknock naphtha rich in normalparailln hydrocarbons which process comprises contacting at least a partof said low knock rating naphtha with a solid, silicaalumina crackingcatalyst in a reaction zone at `a temperature above 575 F. but below 850F.

under an absolute pressure within the approximate range of 15 to 65pounds per square inch and at such space velocity within the approximaterange of about .0l to 1 volume of liquid feed per hour per volume ofcatalyst space in the reaction zone that the conversion of normalparafns per pass to normally gaseous products is only about 5% to 15%,fractionating the products to obtain an isobutane fraction and anormally liquid fraction rich in parafnic hydrocarbons and of higheroctane number than the original charge, recycling a substantial portionof the normally liquid fraction to said reaction zone and removing theremainder 'of said normally liquid

